Polymerisation process

ABSTRACT

A process for the polymerisation of olefins is disclosed wherein at least part of a stream, preferably a catalytically active stream, withdrawn from a polymerisation reactor is passed through a fractionator so as to remove hydrogen and active fines.

The present invention relates to a process for the treatment of areactive fluid stream, and more particularly to a process for treating astream of gas and/or liquid containing active polymer withdrawn from apolymerisation reactor so as to remove reagents and/or impuritiestherefrom. Options available to treat reactive fluid streams can beconstrained by the potential for fouling, blockages etc caused by solidspolymerized or carried in the fluid stream. The present invention aimsto provide an improved process for treating diluent streams in theproduction of polyolefins. Such processes are useful where it isrequired to recycle or feed to a reactor a fluid stream having a lowerlevel of reagent and/or impurities than the level when the stream waswithdrawn from the same or a different reactor.

An example is in multimodal polymerisation reactions, where the polymersare typically manufactured in reactors connected in series with reactionconditions being differentiated in each reactor. In order to havemaximum control of final product properties, it is preferred to havefull and independent control of the molecular weight and the density ofthe polymer produced in each reactor; molecular weight is typicallycontrolled using hydrogen. Consequently it is usually necessary toremove hydrogen from the product stream of a first reactor operating athigher hydrogen concentration than a second reactor in series with thefirst. The polymer from an upstream reactor to be fed to a downstreamreactor in series is typically withdrawn with diluents (gaseous and/orliquid), catalysts and reagents such as monomer(s), comonomer(s),molecular weight control agents such as hydrogen, and cocatalysts.Various technical solutions are known to remove these undesired diluentsand/or reagents, including hydrogen, either fully or partially from thepolymer prior to its entry into the downstream reactor(s). Suchtechniques typically include pressure reduction to vaporise undesiredcomponents.

In EP 603935A a process is described in which a bimodal polyethylene isproduced in series reactors with a low molecular weight homopolymercomponent being formed in the first reactor and a high molecular weightcopolymer component being incorporated in the second reactor, withhydrogen being used to control the molecular weight. There is nodiscussion of how to remove residual hydrogen between the reactors. InEP 192427A and EP 897934A a significant pressure reduction between thetwo reactors is employed to remove at least a portion of the hydrogenpresent. This process is acceptable when the diluent remainssubstantially in the liquid phase under the pressure reductionconditions required to achieve the desired removal of hydrogen: howeverif a more volatile diluent is employed, or if a greater degree ofhydrogen separation is required, then a more effective method isdesirable. Slurry processes employing light (ie relatively volatile)solvents exhibit certain advantages over heavier solvent systems. Forexample, polyolefin oligomers tend to be less soluble, and the solventis readily and substantially completely removed from the polymerproduct. However, hydrogen gas must be virtually completely removedbetween prior and subsequent stages, otherwise process control of thesubsequent stage is difficult and high molecular weight may beimpossible to attain. Light solvents tend to flash away with thehydrogen. If too much solvent flashes off, solids in the slurry take-offincrease to such a high level that the slurry may no longer be pumpable.If solvent flash-off is reduced, hydrogen separation is poor. A furtherdifficulty is that polymer entrained in the flash gas is stillcatalytically active and may polymerise further, causing problems withfouling of any apparatus employed for hydrogen removal or otherseparation. Thus it is necessary either to remove or to deactivate theresidual polymer. Hence it can be seen that for this type ofpolymerisation reaction there is a need for an improved process forremoving hydrogen between stages.

In US 2003/0191251 two flash vessels are used for the separation ofhydrogen from a light diluent between polymerisation reactors. Eachvessel has only one equilibrium stage. Significant diluent make-up isrequired after the first flash step due to a high diluent loss.

In U.S. Pat. No. 3,658,780 a polypropylene slurry withdrawn from apolymerisation reactor is treated with catalyst removal agents and thecatalyst then washed out, thereby rendering the stream catalyticallyinactive, prior to fractionation of the stream to remove hydrogen.

In U.S. Pat. No. 6,045,661 a stream withdrawn from a reactorpolymerising ethylene and hexene in isobutane is passed through flashvessels, and entrained polymer particles removed in a cyclone. At leasta portion of the vapour is then compressed before being passed to afractionator to separate the components. In this process it is statedthat removal of entrained solids ensures that the fractionated materialis not catalytically active.

The present invention aims to provide an improved process for treatingstreams of polyolefins, particularly polyethylene.

Accordingly, in a first aspect the present invention provides a processfor the polymerisation of olefins wherein at least part of a streamwithdrawn from a polymerisation reactor is fractionated in afractionator which comprises a column having at least 3 equilibriumstages. Preferably the process is a continuous polymerisation process.

Typically a stream withdrawn from a polymerisation reactor contains atleast 0.005 particle vol % solid polymer. In such a case the polymercontains the active catalyst. Such solid polymer usually has a particlesize such that at least 50% of the polymer has a particle size of atleast 10 μm. In one embodiment the polymer concentration consists offine particles having a mean diameter of less than 100 microns,preferably less than 50 microns. In another embodiment of the inventionthe concentration of polymer fed into the fractionation stage is atleast 30 vol %, and may be higher than 40 vol %.

Typically the stream withdrawn from the polymerisation vessel iscatalytically active, by which is meant that the stream is capable ofundergoing further polymerisation under the conditions present duringfractionation.

In this specification fractionation means separation in a vessel (i)having more than one equilibrium stage (ii) in which liquid and gas comeat least partially into contact in each equilibrium stage, and (iii)within which the fluid stream is vapourised more than once, preferablymore than twice. “Fractionator” means a vessel or column in whichfractionation takes place.

Reference to an equilibrium stage shall mean a real contacting stage asopposed to a theoretical equilibrium stage.

The fractionation of the stream is preferably carried out in afractionator at a pressure lower than that in the precedingpolymerisation reactor(s) and such that the principal fluid in thestream is condensable without recompression, by heat exchange with acooling medium in the temperature range of about 15-60° C. Mostpreferably the fractionation is carried out at a pressure andtemperature such that at least 50 wt %, preferably at least 75 wt % ofthe catalytically active fluid stream (absent solid component) which isfed to the fractionator is in the vapour phase.

In this specification, “diluent” means a hydrocarbon component added tothe polymerisation reactor to assist removal of heat and/or tosuspension of the solid polymer in the reactor. In the case of slurryreactors the diluent is in the liquid or supercritical state in thereactor. The principal diluent is the non-solid component of the fluidstream having the greatest mol % within the reactor, and is preferablyinert (ie does not polymerise) under reaction conditions.

The treatment of a stream of polymer which may also contain monomer(s)would be expected to cause unacceptable fouling and/or equipmentdowntime to a continuous polymerisation process, particularly to theinternals of, and/or the heat exchange equipment associated with, afractionation column. It has however been found that the process of theinvention may be operated without undue fouling or downtime and that theinstallation of stand-by treatment facilities can be avoided. Anadvantage of separating undesired light components using fractionationat medium to high pressure with more than one equilibrium stage, ratherthan a lower pressure single flash drum, is that less recompression ofthe polymer stream and/or recovered light materials is required, and amore efficient separation operation is also possible. Particularly whenseparating undesired light components from desired light diluents, theloss of diluent from the separation process is significantly reduced.

The process of the invention includes within its scope the treatment ofpolymer streams withdrawn from more than one reactor. In a preferredembodiment of the invention a fractionator is able simultaneously totreat diluent streams from a combination of reactors (in series orparallel) and return diluent to those reactors that is essentially freeof light and/or heavy components, for example streams that are free ofhydrogen, or free of comonomer, or free of the principal monomer, orfree of all monomers. Having the ability to withdraw treated diluentfrom any equilibrium stage within the fractionator gives significantflexibility to optimise the quantity and purity of each recycle streamin an economic manner.

This invention is particularly suited to polymerisation of olefins inslurry or suspension reactors. In this case olefin(s) are continuouslyadded to a hydrocarbon diluent (which may be principally inert orprincipally monomer and be principally a liquid or a supercriticalfluid) containing catalyst. The monomer(s) polymerise to form a slurryof solid particulate polymer suspended in the polymerisation medium ordiluent. Typically, in the particle-form or slurry polymerisationprocesses for polyethylene, the composition of the slurry in the reactoris: particulate polymer about 15-50 particle vol %, preferably 25-40particle vol %; suspension fluid of about 30-85 vol %, and monomer ofabout 1-15 vol %, where the principal diluent is an inert diluent,though these proportions can vary considerably. Particle volume % (vol%) is defined as the volume of particles (excluding interstices volume)in the mixture, divided by the total volume of the mixture.

The invention is most preferably related to polymerisation in anelongated tubular closed loop reaction zone or the so-called ‘slurryloop’ reactor. In the slurry loop process, the reactor is a tubular loopof eg steel pipe located within a larger pipe through which water flowsto heat or cool the reactor as desired. One or more circulating pumpsdrive the reactor contents around the loop at relatively high velocity,in order to promote good heat transfer, maintain the solids insuspension, and to minimize reactor fouling. The loops may be orientedhorizontally or vertically. Product take-off from a loop reactor may becontinuous, or via periodically opened settling legs. In both cases,solvent slurry medium is removed along with the product, and must becondensed and/or repressurised and reintroduced into the reactor.

In a typical polymerisation process to which the present invention isparticularly applicable, the homopolymerisation and copolymerisationprocesses are carried out in the liquid phase in an inert diluent, andthe reactants comprise ethylene and hydrogen in the case ofhomopolymerisation, and in the case of copolymerisation they compriseethylene, alpha-olefinic comonomer(s) comprising from 3 to 8 carbonatoms and optionally hydrogen. The comonomer may be selected frompropylene, 1-butene, 1-pentene, 1-hexene, 4-methyl 1-pentene, 1-hepteneand 1-octene. The inert diluent may comprise (iso)butane, pentane orhexane. The homopolymerisation and copolymerisation processes aretypically carried out at a temperature of from 50 to 120° C., at anabsolute pressure of 1 to 100 bar.

Such a process can be used to make multi-modal polymers, either in asingle reactor or in multiple reactors connected in series or inparallel, in which case any reactor may be preceded or followed by areactor that is of the same or different type of reactor (eg gas phase,stirred slurry tank or loop reactor or solution reactor). In the case ofseries reactors, a first reactor of the series is supplied with acatalyst, and optionally a cocatalyst, and each subsequent reactor issupplied with, at least, ethylene and with the slurry arising from thepreceding reactor of the series, this mixture comprising a catalyst,optionally a cocatalyst and a mixture of the polymers produced in thepreceding reactors of the series. It is possible to supply to a secondreactor and/or, if appropriate, at least one of the following reactorsfresh catalyst and/or cocatalyst, although typically the catalyst andthe cocatalyst are introduced exclusively into the first reactor.

The process of the present invention is especially suitable for apolymerisation which comprises the use of at least two slurry reactorsin series to produce a polyolefin product where a subsequent slurryreactor employs little or no hydrogen feed compared to the hydrogen feedto a prior slurry reactor. In this case the process of the invention isutilised to remove hydrogen from the intermediate polymer slurry betweenthe cascaded reactors. Usually two slurry reactors are employed,although it is also possible to employ three or more reactors in series.It is also possible to employ two or more slurry reactors in seriesalong with one or more slurry reactors operating concurrently inparallel. In this preferred operation of the present invention, afractionator is employed to remove hydrogen and also separate othercomponents of the stream.

In a preferred embodiment of the invention, the polymer stream withdrawnfrom the polymerisation reactor is treated prior to fractionation so asto either minimise the quantity of diluent to be treated and/or tocontrol the average particle size and particle size distribution of thecatalytically active material in the diluent stream. The polymer streamis preferably concentrated to achieve solids concentrations of 50-70 wt%. Preferably the treatment comprises feeding the polymer stream to ahydrocyclone separator prior to fractionation, most preferably with afresh diluent feed upstream of the hydrocyclone separator, for exampleas described in our patent number EP 1118624A. Alternatively asufficiently high solids concentration may be achieved by the use ofsettling legs in the reactor.

For the purposes of the present invention, the term “hydrocycloneseparator” is intended to denote any apparatus which, under the actionof a centrifugal force, makes it possible to separate from a suspensionof solid particles, on the one hand a liquid flow depleted in solidparticles, and on the other hand a flow concentrated in solid particles.Such items of apparatus are well known and are described in particularin Perry's Chemical Engineers' Handbook, McGraw-Hill 7th Edition, 1997,pages 19-24 to 19-28.

If the solids loading of the polymer stream is maintained at asufficiently high concentration (through appropriate operation of thereactor and/or above-mentioned solids concentration system), the heatcontent of the polymer stream may be sufficient to provide all the heatnecessary for fractionation. In such a case, the stream can be feddirectly from the reactor or concentrator into the fractionator forseparation without the addition of any further heat. However if the heatcontent of the solids entering the fractionator is insufficient toprovide all of the necessary reboil for the fractionation column, thebase of the fractionator may be heated, using either a jacket or a heatexchanger. Preferably, the heat content of the polymer stream issufficient to provide at least 60%, more preferably at least 70%, of theheat necessary for fractionation.

In an alternative embodiment, the polymer stream withdrawn from thereactor is heated, optionally after concentration and/ordepressurisation, in the withdrawal line. The resultant stream of solid,gas and optionally liquid is then fed directly to a fractionation columnfeed vessel or alternatively directly to the fractionation column.

If a fractionator feed vessel is used, it is preferred in one embodimentthat the pressure therein is adjusted so as to flash off sufficientdiluent from the feed vessel and leave an unsuspended polymer in thebase. The flashed diluent stream is then fed to the fractionator vessel,the base of which may be heated if necessary. Other than the use of theslurry line heaters and the column base or feed vessel heating jacket,it is preferred that no external energy is added as reboil for thefractionator column.

Where a suspension is maintained in the fractionator feed vessel thepressure is dictated by the need to minimize the concentration ofundesired components in the liquid phase whilst also being able tocondense the principal diluent without compression using only coolingmedium, preferably water, at a temperature between 15 and 60° C. Theflash may be assisted by heating in the feed vessel if necessary. Theflashed diluent stream may be fed directly to the lower part of thefractionation column, preferably the base. In this case the column isdesigned to handle a large particle size and flow-rate below the feedpoint and to accommodate fines in the separation stages immediatelyabove the feed point. In this case, the fractionation column preferablyhas at least 4 actual equilibrium stages, and it is preferred that atleast two separation stages above the feed location of the polymerstream are designed to handle solids. The stripping liquid mass flow inthis part of the fractionation column, preferably throughout thefractionator, is preferably at least 10wt % of the vapour mass flowrate. The fractionator may have internal components such as distillationtrays (sieve, dual-flow, bubble cap, doughnuts) or may equally bepacked, preferably with larger opening packings. The portion of thefractionation column below the feed point is particularly designed toaccommodate a high solids concentration and to avoid solids buildup. Thebase of the column is designed to ensure a high slurry velocity to avoidsedimentation and minimize residence time. The residence time of anysolids in the column is preferably maintained at less than 30 seconds,preferably no more than 90 seconds.

In the fractionation column, liquid bottoms product, containing polymer,diluent and heavier hydrocarbons and/or comonomers, is withdrawn fromthe base of the column. This may optionally be recycled to thefractionator feed vessel. Overhead vapour from the column typicallycontains hydrogen and monomer. Sidedraw streams may also be taken off,containing varying compositions of diluent and monomer. These may berecycled to the reactor.

The fractionation column base temperature should always be maintained atleast 5° C. below the sintering temperature or solubility temperature ofthe polymer produced in an upstream reactor. The design and reliableoperation of the separation process is optimised through carefulequipment design and selection and control of the fluid flowrates andtemperature and pressure conditions through the fractionation column toensure that the solids do not collect at any point in the fractionatorand that fouling of the internals is avoided or at least minimized tosuch an extent that any required cleaning operations do not inthemselves reduce plant availability. Typically the fractionator designis optimized with a base temperature of above 50° C. In one preferredembodiment of the invention the profile of the column temperature andconcentration of the principal monomer concentration profile is designedto avoid excessive peaks in the catalytic activity in any particularstage of the column.

The fractionation is preferably carried out a pressure lower than thatin the preceding reactor but greater than 1 barg, preferably greaterthan 3.5 barg.

The slurry in the base of the fractionator column, and in certainembodiments in the base of the fractionator feed vessel, is preferablykept in suspension at all times, and is usually agitated with a stirrer;however it may alternatively or additionally be suspended using anexternal circulation pump.

In the case where unsuspended polymer powder is maintained in thefractionator feed vessel, the base may be designed to enable a constantpowder level to be maintained whilst allowing continuous ordiscontinuous flow out of the vessel. The base is preferably conical inshape with the angle of the cone and the outlet nozzle diameter sized soas to maintain plug or mass flow for the range of powders (taking intoaccount the expected associated hydrocarbon content) that the vessel isdesigned to handle.

In a preferred embodiment, the fractionator base is discharged to aslurry vessel that receives both slurry feed from the fractionator baseand also suspended or unsuspended solids from the fractionator feedvessel base. In this case the solids in the fractionator feed vessel arepreferably unsuspended. Fresh diluent may be added to the transfer linebetween the fractionator feed vessel base and the slurry vessel. In thecase where the fractionator is separating a feed stream intermediate twopolymerisation reactors in a bimodal polymerisation, the slurry vesseldischarges to the second reactor.

A specific example of a process for application of the present inventionis a suspension polymerisation process for the production of a bimodalhigh density polyethylene comprising an ethylene homopolymer (A) and acopolymer of ethylene and 1-hexene (B) formed in two reactors in series.Such a process may be carried out using the apparatus shown in FIG. 2(see below).

The diluent used in this particular polymerisation process is usually ahydrocarbon-comprising diluent which is inert with respect to thecatalyst, the cocatalyst and the polymer formed, such as, for example, alinear or branched alkane or a cycloalkane having from 3 to 8 carbonatoms. The diluent which has given the best results is isobutane. Oneadvantage of the use of isobutane lies in particular in its readyrecycling. This is because the use of isobutane makes it possible torecycle the diluent recovered at the end of the process according to theinvention in the first reactor without having to carry out exhaustivepurification in order to remove the residual hexene. This is because, asthe boiling temperatures of isobutane and of hexene are far apart, theirseparation can be carried out by distillation.

In this preferred process, the amount of ethylene introduced into thefirst polymerisation reactor and into the subsequent polymerisationreactor is generally adjusted so as to obtain a concentration ofethylene in the diluent of 5 to 50 g of ethylene per kg of diluent. Theamount of hydrogen introduced into the first reactor is generallyadjusted so as to obtain, in the diluent, a molar ratio of hydrogen toethylene of 0.05 to 1. A hydrogen/ethylene molar ratio which does notexceed 0.6 is particularly preferred.

The mixture withdrawn from the first reactor, additionally comprisinghomopolymer (A), is subjected to a reduction in pressure so as to remove(degas) at least a portion of the hydrogen, which can be conducted inaccordance with the present invention. The reduction in pressure isadvantageously carried out at a temperature of less than or equal to thepolymerisation temperature in the first reactor. The temperature atwhich the reduction in pressure is carried out is usually at least 40°C. The pressure at which the reduction in pressure is carried out isless than the pressure in the first reactor, and is usually between 0.1and 1.5 MPa. The amount of hydrogen still present in the at leastpartially degassed slurry (liquid+solid) mixture is generally less than1% by weight of the amount of hydrogen initially present in the mixturewithdrawn from the first polymerisation reactor; this amount ispreferably less than 0.5%. The amount of hydrogen present in thepartially degassed mixture introduced into the subsequent polymerisationreactor is consequently low, or even zero. The subsequent reactor ispreferably also fed with hydrogen. The amount of hydrogen introducedinto the subsequent reactor is generally adjusted so as to obtain, inthe diluent, a molar ratio of hydrogen to ethylene of 0.001 to 0.1 inthe reactor, typically between 0.004 and 0.05. In this process, theratio of the concentration of hydrogen in the diluent in the firstreactor to the concentration in the subsequent polymerisation reactor isusually at least 20, preferably between 40 and 200.

The amount of 1-hexene introduced into the subsequent polymerisationreactor is such that, in this reactor, the hexene/ethylene molar ratioin the diluent is at least 0.05, preferably at least 0.1. The amount ofhexene introduced into the subsequent reactor is such that thehexene/ethylene molar ratio is preferred not to exceed 3. The firstreactor is usually not fed with hexene; indeed, it is essential that thefirst reactor is essentially devoid of 1-hexene. Consequently, thediluent introduced into the first reactor, which can be recycleddiluent, must be highly depleted in hexene. The diluent introduced intothe first reactor preferably contains less than 1000 ppm of hexene, andis ideally essentially free of hexene.

The polymerisation temperature is generally from 20 to 130° C.,typically not exceeding 115° C. The total pressure at which the processaccording to the invention is carried out is generally from 0.1 MPa to10 MPa. In the first polymerisation reactor, the total pressure isusually at least 2.5 MPa, but not greater than 5 MPa. In the subsequentpolymerisation reactor, the total pressure is usually at least 1.3 MPa,but not greater than 4.3 MPa.

In this preferred process, a suspension comprising a compositioncomprising from 30 to 70% by weight of the homopolymer (A) and from 30to 70% by weight of the copolymer (B) is collected at the outlet of thesubsequent polymerisation reactor. The composition comprising ethylenepolymers can be separated from the suspension by any known means. Thesuspension is usually subjected to a reduction in pressure (finalreduction in pressure), so as to remove the diluent, the ethylene, thehexene and, optionally, the hydrogen from the composition.

According to an alternative form of this process and more particularlywhen the diluent is isobutane, the gases exiting from the firstreduction in pressure (intermediate reduction in pressure between thetwo polymerisation reactors) and from the final reduction in pressureare mixed, and conveyed to a distillation unit. This distillation unitis advantageously composed of one or of two distillation columns inseries. Ethylene and hydrogen are withdrawn at the column top, a mixtureof isobutane and of hexene is withdrawn at the column bottom andisobutane devoid of hexene is withdrawn from an intermediate plate. Theisobutane-hexene mixture is then recycled in the subsequentpolymerisation reactor, whereas the isobutane devoid of hexene isrecycled in the first reactor.

The catalyst employed in the polymerisation process may be anycatalyst(s) suitable for polymerisation reactions, but is typically achromium catalyst, a Ziegler-Natta catalyst, or a metallocene catalyst.Usually the catalyst is a Ziegler-Natta catalyst.

In the case of a Ziegler-Natta catalyst, the catalyst used comprises atleast one transition metal. Transition metal means a metal of groups 4,5 or 6 of the Periodic Table of elements (CRC Handbook of Chemistry andPhysics, 75th edition, 1994-95). The transition metal is preferablytitanium and/or zirconium. A catalyst comprising not only the transitionmetal but also magnesium is preferably utilised. Good results have beenobtained with catalysts comprising:

-   -   from 10 to 30%, preferably from 15 to 20%, more preferably 16 to        18% by weight of transition metal,    -   from 0.5 to 20%, preferably from 1 to 10%, more preferably 4 to        5% by weight of magnesium,    -   from 20 to 60%, preferably from 30 to 50%, more preferably 40 to        45% by weight of halogen, such as chlorine,    -   from 0.1 to 10%, preferably from 0.5 to 5%, more preferably 2 to        3% by weight of aluminium;

the balance generally consisting of elements arising from products usedfor their manufacture, such as carbon, hydrogen and oxygen. Thesecatalysts are preferably obtained by coprecipitation of at least onetransition metal composition and a magnesium composition by means of ahalogenated organoaluminium composition. Such catalysts are known, theyhave notably been described in patents U.S. Pat. No. 3,901,863, U.S.Pat. No. 42,942,200 and U.S. Pat. No. 4,617,360. The catalyst ispreferably introduced only into the first polymerisation reactor, i.e.there is no introduction of fresh catalyst into the furtherpolymerisation reactor.

The cocatalyst utilised in the process is preferably an organoaluminiumcompound. Unhalogenated organoaluminium compounds of formula AlR₃ inwhich R represents an alkyl grouping having from 1 to 8 carbon atoms arepreferred. Particularly preferred are triethylaluminium andtriisobutylaluminium. The cocatalyst is introduced into the firstpolymerisation reactor. Fresh cocatalyst may also be introduced into thefurther reactor. The quantity of cocatalyst introduced into the firstreactor is in general at least 0.1×10⁻³ mole per litre of diluent. Itdoes not usually exceed 5×10⁻³ mole per litre of diluent. Any quantityof fresh cocatalyst introduced into the further reactor does not usuallyexceed 5×10⁻³ mole per litre of diluent.

Specific embodiments of the present invention will now be described withreference to the accompanying drawings.

FIG. 1 shows a flowsheet for a polymerisation system comprising a singleslurry loop polymerisation reactor with associated fractionator.

Referring to FIG. 1, a diluent is maintained in a liquid phase during apolymerisation reaction where the polymer solids produced areessentially not soluble in the diluent and are suspended by it. Theeffluent stream of the polymerisation reactor 1 comprises a liquiddiluent carrying a slurry of polymer solids together with residualcatalyst and reagents such as monomer(s), comonomer(s), molecular weightcontrol agents such as hydrogen, and cocatalysts.

The effluent stream is withdrawn from the reactor via line 3 from whereit passes into a hydrocyclone 5 which concentrates the slurry to asolids level of about 50-70 wt %. The stream is then usually subjectedto a pressure let-down at 7, from the reactor pressure (typically 40barg) to a pressure of 7-10 barg. Depending on the solids concentrationand temperature of the stream, the heat content of the stream may beboosted by slurry line heater 9; the degree of heat input designed orcontrolled to maximise vaporization of liquid whilst avoiding risk ofsintering in the heater. Preferably the slurry heater exit temperatureis controlled to the dew point temperature of the stream of fluidwithdrawn.

The stream then passes into a fractionator feed vessel 11. The pressurein the feed vessel 11 is adjusted so as to flash off sufficient diluentto leave an unsuspended polymer in the base. The solid polymer iswithdrawn through line 13.

The flashed diluent stream is then fed via line 15 towards thefractionator column 17, preferably at the base. The base of the columnmay be heated if the solids content of the stream is insufficient toprovide enough heat for fractionation. The pressure at which column 17operates can be in the broad range of 1 barg to 30 barg or more.Preferred temperature conditions in column 17 include an overheadtemperature (temperature at the top of the column) of 30-50° C. and abottoms temperature (temperature at the bottom of the column) of 65-95°C. The fractionator has between 5 and 25 sieve and/or dual flow trays.

Liquid bottoms product, typically containing diluent rich in heavycomonomer(s), is withdrawn from column 17 through line 19. If additionalheating is required, some of the bottoms product is passed through line19 to a heater (reboiler) 21, and from there through line 23 as vapourback to column 17. Alternately, the bottom of Column 17 can be heatedwith a jacket.

In the case where the comonomer is heavier than the diluent, a(preferably vaporous) sidedraw stream may optionally be withdrawn fromcolumn 17 through line 25. The sidedraw stream typically contains mainlydiluent lean in comonomer. The sidedraw stream is cooled and condensedand then recycled to the reactor 1 (not shown). It can also be withdrawnfrom the column as a liquid. The column, by providing a stream leaner incomonomer than in the preceding reactor, and or by providing buffercapacity of comonomer lean streams provides the facility tosubstantially reduce the time for product transitions between polymergrades of differing density.

Overhead vapour from column 17, typically containing diluent, unreactedmonomer, hydrogen, nitrogen, and other lights passes through line 27 tocooler 29, where it is condensed to be recycled as reflux to the column17 via line 31. A lights vent may be taken from the condenser.

Referring to FIG. 2, this shows an alternative embodiment of theinvention relating to a bimodal polymerisation in which hydrogen isremoved from the polymerisation stream by a fractionator located betweenthe reactors. The second reactor is not shown in the Figure. In thearrangement shown in FIG. 2, the numerals are the same as in FIG. 1. Theprincipal difference from the arrangement of FIG. 1 is that in thisembodiment the bottoms product from the fractionator 17 is recycled tothe fractionator feed vessel 11 via line 33. In this case the polymer inthe bottom of the feed vessel 11 is maintained in suspension in thediluent by a stirrer 35, and this suspension is withdrawn from the baseof the feed vessel via line 13 and pumped by pump 37 to the secondreactor (not shown). The liquid portion of the stream withdrawn throughline 13 may be recycled to the feed vessel 11 via line 39, which maycontain a heater 40. The feed vessel 11, having a well agitated solidssuspension, may be heated by a heater 43.

Regarding the fractionator 17 in the embodiment of FIG. 2, this operatesas in FIG. 1 without any sidedraws except that hydrogen is vented fromthe overhead stream, line 27, when the stream is otherwise condensed inthe cooler 29. Thus the recycled stream 31 has a substantially reducedlevel of hydrogen. Stream 33 recovers the vast majority of diluent,comonomers and even monomer flashed in vessel 11 whilst being lean inhydrogen. This fractionator design typically has about 5 sieve and/ordual flow trays.

Referring to FIG. 3, this shows a further embodiment of the inventionrelating to a bimodal polymerisation in two reactors, in which a singlefractionation column is used to treat both the polymer streamintermediate the reactors and also the final stream from the secondreactor. In the following description it is assumed that the lowmolecular weight product is made in the first reactor, however thisdesign configuration gives full flexibility, by choosing the appropriaterecycle streams from the fractionator, to equally enable the highmolecular weight product to be made in the first reactor.

As in the embodiments of FIGS. 1 and 2, effluent stream from the firstreactor 1 is withdrawn from the reactor via line 3 from where it passesinto a fractionator feed vessel 11 (details shown in FIGS. 1 and 2omitted here). As in the previous embodiments, concentration with ahydrocyclone, a pressure let-down and additional heating (none shown inthe Figure) may all be applied to the stream if required.

Diluent is flashed from the feed vessel 11 via line 15 to afractionation column 17. All or most of comonomer rich diluent streamfrom column 17 is returned to vessel 11 via line 57. Suspended polymeris withdrawn from the bottom of the feed vessel 11 via line 13 and istransferred to the second reactor 41, where additional comonomer may beadded as desired.

Effluent stream from the second reactor 41 is withdrawn from the reactorvia line 43 from where it passes into a second fractionator feed vessel51. As in the case of the first reactor 1, concentration with ahydrocyclone, a pressure let-down and additional heating (none shown inthe Figure) may all be applied to this second effluent stream ifrequired. A powder level is maintained in second fractionator feedvessel 51 and the unsuspended, final polymer product essentially absentof free liquid, is withdrawn via line 55, whilst the vapourised diluentstream is then fed via line 53 to the fractionation column 17.

The temperature profile and pressure of column 17 is adapted so as toremove separate the full spectrum of components of the incoming stream.Preferred temperature conditions in column 17 include an overheadtemperature (temperature at the top of the column) of 35-55° C. and abottoms temperature (temperature at the bottom of the column) of 65-95°C. Liquid bottoms product, together with any polymer fines, typicallycontaining diluent and rich in comonomer, is withdrawn from column 17through line 57. Optional sidedraw stream 61 takes off diluent free ofmonomer if required for example for transport of catalyst to reactor 1.Sidedraw stream 63 consists of a diluent stream free of comonomer thatmay be recycled to reactor 1. Overhead vapour from column 17 withdrawnin line 27 and a stream rich in hydrogen, together with some monomer isvented from condenser 29 through line 65. The number of trays requiredin the column is minimised by preferably not designing it to separatemonomer from the diluent streams recycled to either reactor.

Referring to FIG. 4, this shows an embodiment of the invention relatingto a bimodal polymerisation in which hydrogen is removed from thepolymerisation stream by a fractionator located between the reactors.Thus this is similar to the embodiment of FIG. 2, and where appropriatereference numerals are the same. The effluent stream passes from reactor1 to fractionator feed vessel 11 in the same manner as in the embodimentof FIG. 2. The pressure in the feed vessel 11 is adjusted so as to flashoff sufficient diluent to the fractionator 17 via line 15 to leave anunsuspended polymer in the base. Line 15 enters the fractionator 17above the bottom so as to enhance hydrogen separation. The solid polymeris withdrawn through line 67 into a second slurry vessel 69, which alsoreceives the bottoms product from the fractionator 17 via line 71. Areboiler can also be present at the bottom of column 17 to enhanceseparation in the column. The polymer in the bottom of the slurry vessel69 is in suspension in the hydrogen lean diluent, and this suspension iswithdrawn from the base of the slurry vessel 69 via line 73 and pumpedby pump 37 to the second reactor 41. A liquid portion of the streamwithdrawn through line 73 may be recycled to the slurry vessel 69 vialine 39.

EXAMPLE

In one specific example of the invention in which a fractionator isemployed between two reactors in a bimodal dual-reactor system such asdescribed in FIG. 2, the fractionator is fed with a stream comprisingisobutane, ethylene, hydrogen, hexene-1 and polyethylene issuing from afirst polymerisation reactor. This stream is first concentrated in ahydrocyclone, after which it is passed through a slurry heater beforeentering the base of the fractionator.

In this particular example, a flow rate of 10090 kg/h of polyethyleneand 9685 kg/h hydrocarbon enters the base of the fractionator. Thehydrocarbon in this case comprises mainly isobutane but also containsabout 3.22 kg/h of hydrogen, 116 kg/h ethylene, about 10 kg/h of solidpolymer which contains some active catalyst, and minor amounts of othercomponents.

As the stream enters the base of the fractionator, the hydrocarbonportion is about 75% vapour. Upon entering the base, the residual liquidand almost all of the solid polyethylene falls into a boiling andagitated pool. The base slurry in this example is heated by a jacketwhich provides the heat for about a third of the column's vapour flow,and is at about 70° C. and 10 barg pressure.

The hydrocarbon vapour coming from the reactor stream combines with theboil-up from the boiling slurry in the fractionator base such that avapour flow of about 11500 kg/hr enters the fractionation column. Thisvapour contains some catalytically active polymer fines which carry overfrom the upstream equipment. The fractionation column is about 1 metrein diameter, and has five dual-flow trays. Each tray has about 9% openarea, and has holes of 25 mm diameter. This large hole diameter isimportant to ensure that blocking is minimised.

The gas stream works its way up the scrubber, and each trayprogressively removes fines by contacting the gas/fines stream with aliquid stream which is falling down the column. This also removeshydrogen from the liquid which falls down the column.

At the top of the column, the stream enters a condenser and is almosttotally condensed at about 30° C. A gas purge is taken off thiscondensed stream to remove hydrogen. A filter may be installed in thisstream to test the solids removal efficiency of the column: theApplicants have never found any trace of polyethylene in such a filter.Additionally, the Applicants have never found any polyethylene in thecondenser, nor have they experienced any sort of fouling. Theseobservations confirm the excellent performance of the system in handlingactive fines.

The liquid which is condensed is returned to the fractionator and fallsback down each tray. By the time it reaches the bottom of thefractionator the liquid has been depleted of hydrogen, and any activefines are recycled back to the base liquid. It should be noted thatsince almost all the liquid is condensed and returned back to thefractionator base, there is no need for liquid make-up in the base tomaintain solids concentration. The essentially hydrogen-free slurry isagitated to minimise settling, and is then pumped to the second reactor.A slip stream is taken off this pump discharge and returned to theslurry base to aid in slurry homogeneity. The typical hydrogen contentof the slurry going to the second reactor is below 100 g/h. Thus,considering the initial hydrogen flow rate leaving the first reactor of3.22 kg/h, it can be seen that the process of the invention is veryefficient at removing hydrogen from the feed stream and also removingactive fines from the purge gas at the top of the column. This simpleequipment thereby demonstrates a reliable and economic means to controlhydrogen concentration and hence molecular weight in the second reactorindependently of conditions required in the first reactor—even whenusing diluents which are vapour under atmospheric conditions.Furthermore, the column may also be adjusted so as to minimise theamount of ethylene and diluent lost.

1. Process for the polymerisation of olefins wherein at least part of astream withdrawn from a polymerisation reactor is passed through afractionator which comprises a column having at least 3 equilibriumstages.
 2. Process according to claim 1, which is a continuous process,in which the stream is preferably continuously withdrawn.
 3. Processaccording to claim 1, wherein the stream leaving the polymerisationreactor contains at least 10 vol % solid polymer, preferably greaterthan 30 vol % and more preferably greater than 40 vol % solid polymer.4. Process according to claim 3, wherein the solid polymer fed into thefractionator has a particle size such that at least 50% of the polymerhas a particle size of at least 7, preferably at least 10 μm.
 5. Processaccording to claim 3, wherein at least 50% of the solid polymer has aparticle size less than 100 μm, preferably less than 50 μm.
 6. Processaccording to claim 1, wherein the concentration of polymer in the streamfed into the fractionator is at least 0.002 vol %, preferably at least30 vol % and more preferably at least 40 vol %.
 7. Process according toclaim 6, wherein at least 50% of the solid polymer has a particle sizeless than 2000 μm, preferably less than 1000 μm.
 8. Process according toclaim 1, wherein the heat content of the stream entering thefractionator is sufficient to provide at least 60% of the heat necessaryfor fractionation, and preferably all of the heat necessary forfractionation.
 9. Process according to claim 1, wherein the streamwithdrawn from the polymerisation reactor is fractionated at a pressuresuch that the principal diluent of the stream is substantially condensedwithout compression using only cooling medium, preferably water, at atemperature between 15 and 60° C.
 10. Process according to claim 9,wherein there is no reduction in pressure to below fractionationpressure applied to the stream between the reactor and the fractionator.11. Process according to claim 1, wherein the equilibrium stages in thefractionator comprise sieve trays and/or dual flow trays.
 12. Processaccording to claim 1, wherein at least the first two equilibrium stagesin the fractionator above the feed location of the stream have astripping liquid flow that is at least 10 wt % of the vapour flow rate.13. Process according to claim 12, wherein every equilibrium stage inthe fractionator has a stripping liquid flow that is at least 10 wt % ofits vapour flow rate.
 14. Process according to claim 1, wherein theresidence time of any solids in the fractionator is maintained at nomore than 90 seconds, preferably no more than 30 seconds.
 15. Processaccording to claim 1, which process is a process for the polymerisationof ethylene or propylene utilising more than one reactor in series, andthe principal diluent is an inert diluent or monomer, preferably beingpropylene and/or isobutane and/or hexane.
 16. Process according to claim15, wherein hydrogen and/or comonomer is employed in at least one of thepolymerisation reactors, and fractionation removes at least some of thehydrogen and/or comonomer from at least one of the streams containingthe principal diluent prior to its distribution to a downstream reactor.17. Process according to claim 16, wherein both comonomer and hydrogenare at least partly removed from the principal diluent in the samefractionator.
 18. Process according to claim 15 wherein the fractionatoris fed from more than one reactor and/or the fractionator feeds purifiedprincipal diluent to more than one reactor.
 19. Process according toclaim 15, wherein the ethylene is polymerised to form a polymercomprising at least 30 wt % of a low molecular weight component having adensity of at least 0.965 g/cm3 and an MI2 of from 5 to 1000 g/10 min,and at least 30 wt % of a high molecular weight component having adensity of from 0.910 to 0.940 g/cm3 and an MI5 of from 0.01 to 2 g/10min.
 20. Process according to claim 19, wherein the low molecular weightcomponent is made in a reactor upstream of a reactor making the highmolecular weight component.
 21. Process according to claim 1, whereinthe stream withdrawn from the polymerisation reactor is catalyticallyactive.